Process for oligomerizing dilute ethylene

ABSTRACT

The process and apparatus converts ethylene in a dilute ethylene stream that may be derived from an FCC product to heavier hydrocarbons. The catalyst may be an amorphous silica-alumina base with a Group VIII and/or VIB metal. The catalyst is resistant to feed impurities such as hydrogen sulfide, carbon oxides, hydrogen and ammonia. At least 40 wt-% of the ethylene in the dilute ethylene stream can be converted to heavier hydrocarbons.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a Continuation of copending application Ser. No.12/416,029 filed Mar. 31, 2009, the contents of which are herebyincorporated by reference in its entirety.

BACKGROUND OF THE INVENTION

The field of the invention is an apparatus and process for convertingdiluted ethylene in a hydrocarbon stream to heavier hydrocarbons. Theseheavier hydrocarbons may be used as motor fuels.

Dry gas is the common name for the off-gas stream from a fluid catalyticcracking unit that contains all the gases with boiling points lower thanethane. The off-gas stream is compressed to remove as much of the C₃ andC₄ gases as possible. Sulfur is also largely absorbed from the off-gasstream in a scrubber that utilizes an amine absorbent. The remainingstream is known as the FCC dry gas. A typical dry gas stream contains 5to 50 wt-% ethylene, 10 to 20 wt-% ethane, 5 to 20 wt-% hydrogen, 5 to20 wt-% nitrogen, about 0.1 to about 5.0 wt-% of each carbon monoxideand carbon dioxide and less than 0.01 wt-% hydrogen sulfide and ammoniawith the balance being methane.

Currently, the FCC dry gas stream is sent to a burner as fuel gas. AnFCC unit that processes 7,949 kiloliters (50,000 barrels) per day willburn about 181,000 kg (200 tons) of dry gas with about 36,000 kg (40tons) of ethylene as fuel per day. Because a large price differenceexists between fuel gas and motor fuel products or pure ethylene itwould appear economically advantageous to attempt to recover thisethylene. However, the dry gas stream contains impurities that canpoison oligomerization catalyst and is so dilute that ethylene recoveryis not economically justified by gas recovery systems.

The oligomerization of concentrated ethylene streams to liquid productsis a known technology. However, oligomerization typically involves theuse of propylene or butylene particularly from liquefied petroleum gas(LPG) or dehydrogenated feedstocks to make gasoline range olefins.Ethylene is little used as an oligomerization feedstock because of itsmuch lower reactivity.

There is need for utilization of dilute ethylene in refinery streams.

SUMMARY OF THE INVENTION

We have found that ethylene in dilute ethylene streams, such as an FCCdry gas stream, can be catalytically oligomerized to heavierhydrocarbons with a Group VIII and/or Group VIB metal on amorphoussilica-alumina catalyst. The heavier hydrocarbons can be separated andblended in the gasoline and diesel pools. We have found zeoliticcatalysts that are suitable for oligomerization of ethylene quicklydeactivate in the presence of impurities such as carbon oxides, ammoniaand hydrogen sulfide. The impurities do not substantially affect acatalyst comprising a Group VIII and/or VIB metal on amorphoussilica-alumina support. Consequently, dilute ethylene in an FCC dry gasstream can be oligomerized to a liquid fuel product which is easy toseparate from the unconverted gas stream. The unconverted gas can thenbe burned as fuel gas, but with the more valuable ethylene removed asheavier hydrocarbons.

Advantageously, the process and apparatus can enable utilization ofethylene in a dilute stream and in the presence of feed impurities thatcan be catalyst poisons.

Additional features and advantages of the invention will be apparentfrom the description of the invention, the FIGURES and claims providedherein.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic drawing of an FCC unit and an FCC product recoverysystem.

FIG. 2 is a plot of ethylene conversion versus time on stream forExamples 6-8.

FIG. 3 is a plot of ethylene conversion versus time on stream forExample 9.

FIG. 4 is a plot of ethylene conversion versus time on stream fromExample 10.

DETAILED DESCRIPTION

The present invention may be applied to any hydrocarbon streamcontaining ethylene and, preferably, a dilute proportion of ethylene. Asuitable, dilute ethylene stream may typically comprise between about 5and about 50 wt-% ethylene. An FCC dry gas stream is a suitable diluteethylene stream. Other dilute ethylene streams may also be utilized inthe present invention such as coker dry gas streams. Because the presentinvention is particularly suited to FCC dry gas, the subject applicationwill be described with respect to utilizing ethylene from an FCC dry gasstream.

Now turning to FIG. 1, wherein like numerals designate like components,FIG. 1 illustrates a refinery complex 6 that generally includes an FCCunit section 10 and a product recovery section 90. The FCC unit section10 includes a reactor 12 and a catalyst regenerator 14. Processvariables typically include a cracking reaction temperature of 400° to600° C. and a catalyst regeneration temperature of 500° to 900° C. Boththe cracking and regeneration occur at an absolute pressure below 506kPa (72.5 psia).

FIG. 1 shows a typical FCC reactor 12 in which a heavy hydrocarbon feedor raw oil stream in a distributor 16 is contacted with a regeneratedcracking catalyst entering from a regenerated catalyst standpipe 18.This contacting may occur in a narrow riser 20, extending upwardly tothe bottom of a reactor vessel 22. The contacting of feed and catalystis fluidized by gas from a fluidizing line 24. In an embodiment, heatfrom the catalyst vaporizes the hydrocarbon feed or oil, and thehydrocarbon feed is thereafter cracked to lighter molecular weighthydrocarbon products in the presence of the catalyst as both aretransferred up the riser 20 into the reactor vessel 22. Inevitable sidereactions occur in the riser 20 leaving coke deposits on the catalystthat lower catalyst activity. The cracked light hydrocarbon products arethereafter separated from the coked cracking catalyst using cyclonicseparators which may include a primary separator 26 and one or twostages of cyclones 28 in the reactor vessel 22. Gaseous, crackedproducts exit the reactor vessel 22 through a product outlet 31 to line32 for transport to a downstream product recovery section 90. The spentor coked catalyst requires regeneration for further use. Coked crackingcatalyst, after separation from the gaseous product hydrocarbons, fallsinto a stripping section 34 where steam is injected through a nozzle topurge any residual hydrocarbon vapor. After the stripping operation, thecoked catalyst is carried to the catalyst regenerator 14 through a spentcatalyst standpipe 36.

FIG. 1 depicts a regenerator 14 known as a combustor. However, othertypes of regenerators are suitable. In the catalyst regenerator 14, astream of oxygen-containing gas, such as air, is introduced through anair distributor 38 to contact the coked catalyst. Coke is combusted fromthe coked catalyst to provide regenerated catalyst and flue gas. Thecatalyst regeneration process adds a substantial amount of heat to thecatalyst, providing energy to offset the endothermic cracking reactionsoccurring in the reactor riser 20. Catalyst and air flow upwardlytogether along a combustor riser 40 located within the catalystregenerator 14 and, after regeneration, are initially separated bydischarge through a disengager 42. Additional recovery of theregenerated catalyst and flue gas exiting the disengager 42 is achievedusing first and second stage separator cyclones 44, 46, respectivelywithin the catalyst regenerator 14. Catalyst separated from flue gasdispenses through diplegs from cyclones 44, 46 while flue gas relativelylighter in catalyst sequentially exits cyclones 44, 46 and exits theregenerator vessel 14 through flue gas outlet 47 in flue gas line 48.Regenerated catalyst is carried back to the riser 20 through theregenerated catalyst standpipe 18. As a result of the coke burning, theflue gas vapors exiting at the top of the catalyst regenerator 14 inline 48 contain CO, CO₂, N₂ and H₂O, along with smaller amounts of otherspecies. Hot flue gas exits the regenerator 14 through the flue gasoutlet 47 in a line 48 for further processing.

The product recovery section 90 is in downstream communication with theproduct outlet 31. “Downstream communication” means that at least aportion of material flowing to the component in downstream communicationmay operatively flow from the component with which it communicates.“Communication” means that material flow is operatively permittedbetween enumerated components. In the product recovery section 90, thegaseous FCC product in line 32 is directed to a lower section of an FCCmain fractionation column 92. The main column 92 is in downstreamcommunication with the product outlet 31. Several fractions of FCCproduct may be separated and taken from the main column including aheavy slurry oil from the bottoms in line 93, a heavy cycle oil streamin line 94, a light cycle oil in line 95 taken from outlet 95 a and aheavy naphtha stream in line 96 taken from outlet 96 a. Any or all oflines 93-96 may be cooled and pumped back to the main column 92 to coolthe main column typically at a higher location. Gasoline and gaseouslight hydrocarbons are removed in overhead line 97 from the main column92 and condensed before entering a main column receiver 99. The maincolumn receiver 99 is in downstream communication with the productoutlet 31, and the main column 92 is in upstream communication with themain column receiver 99. “Upstream communication” means that at least aportion of the material flowing from the component in upstreamcommunication may operatively flow to the component with which itcommunicates.

An aqueous stream is removed from a boot in the receiver 99. Moreover, acondensed light naphtha stream is removed in line 101 while an overheadstream is removed in line 102. The overhead stream in line 102 containsgaseous light hydrocarbon which may comprise a dilute ethylene stream.The streams in lines 101 and 102 may enter a vapor recovery section 120of the product recovery section 90.

The vapor recovery section 120 is shown to be an absorption basedsystem, but any vapor recovery system may be used including a cold boxsystem. To obtain sufficient separation of light gas components thegaseous stream in line 102 is compressed in compressor 104. More thanone compressor stage may be used, but typically a dual stage compressionis utilized. The compressed light hydrocarbon stream in line 106 isjoined by streams in lines 107 and 108, chilled and delivered to a highpressure receiver 110. An aqueous stream from the receiver 110 may berouted to the main column receiver 99. A gaseous hydrocarbon stream inline 112 comprising the dilute ethylene stream is routed to a primaryabsorber 114 in which it is contacted with unstabilized gasoline fromthe main column receiver 99 in line 101 to effect a separation betweenC₃+ and C₂— hydrocarbons. The primary absorber 114 is in downstreamcommunication with the main column receiver 99. A liquid C₃+ stream inline 107 is returned to line 106 prior to chilling. A primary off-gasstream in line 116 from the primary absorber 114 comprises the diluteethylene stream for purposes of the present invention. However, toconcentrate the ethylene stream further and to recover heaviercomponents line 116 may optionally be directed to a secondary absorber118, where a circulating stream of light cycle oil in line 121 divertedfrom line 95 absorbs most of the remaining C₅+ and some C₃-C₄ materialin the primary off-gas stream. The secondary absorber 118 is indownstream communication with the primary absorber 114. Light cycle oilfrom the bottom of the secondary absorber in line 119 richer in C₃+material is returned to the main column 92 via the pump-around for line95. The overhead of the secondary absorber 118 comprising dry gas ofpredominantly C₂— hydrocarbons with hydrogen sulfide, ammonia, carbonoxides and hydrogen is removed in a secondary off-gas stream in line 122to comprise a dilute ethylene stream.

Liquid from the high pressure receiver 110 in line 124 is sent to astripper 126. Most of the C₂— is removed in the overhead of the stripper126 and returned to line 106 via overhead line 108. A liquid bottomsstream from the stripper 126 is sent to a debutanizer column 130 vialine 128. An overhead stream in line 132 from the debutanizer comprisesC₃-C₄ olefinic product while a bottoms stream in line 134 comprisingstabilized gasoline may be further treated and sent to gasoline storage.

The dilute ethylene stream of the present invention may comprise an FCCdry gas stream comprising between about 5 and about 50 wt-% ethylene andpreferably about 10 to about 30 wt-% ethylene. Methane will typically bethe predominant component in the dilute ethylene stream at aconcentration of between about 25 and about 55 wt-% with ethane beingsubstantially present at typically between about 5 and about 45 wt-%.Between about 1 and about 25 wt-% and typically about 5 to about 20 wt-%of hydrogen and nitrogen each may be present in the dilute ethylenestream. Saturation levels of water may also be present in the diluteethylene stream. If secondary absorber 118 is used, no more than about 5wt-% of C₃+ will be present with typically less than 0.5 wt-% propylene.

Besides hydrogen, other impurities such as hydrogen sulfide, ammonia,carbon oxides and acetylene may also be present in the dilute ethylenestream.

We have found that many impurities in a dry gas ethylene stream canpoison an oligomerization catalyst. Hydrogen and carbon monoxide canreduce the metal sites to inactivity. Carbon dioxide and ammonia canattack acid sites on the catalyst. Hydrogen sulfide can attack metals ona catalyst to produce metal sulfides. Acetylene can polymerize and gumup on the catalyst or equipment.

The secondary off-gas stream in line 122, comprising a dilute ethylenestream may be introduced into an optional amine absorber unit 140 toremove hydrogen sulfide to lower concentrations. A lean aqueous aminesolution, such as comprising monoethanol amine or diethanol amine, isintroduced via line 142 into absorber 140 and is contacted with theflowing secondary off-gas stream to absorb hydrogen sulfide, and a richaqueous amine absorption solution containing hydrogen sulfide is removedfrom absorption zone 140 via line 143 and recovered and perhaps furtherprocessed.

The amine-treated dilute ethylene stream in line 144 may be introducedinto an optional water wash unit 146 to remove residual amine carriedover from the amine absorber 140 and reduce the concentration of ammoniaand carbon dioxide in the dilute ethylene stream in line 144. Water isintroduced to the water wash in line 145. The water in line 145 istypically slightly acidified to enhance capture of basic molecules suchas the amine. An aqueous stream in line 147 rich in amine andpotentially ammonia and carbon dioxide leaves the water wash unit 146and may be further processed.

The optionally amine treated dilute ethylene and perhaps water washedstream in line 148 may then be treated in an optional guard bed 150 toremove one or more of the impurities such as carbon monoxide, hydrogensulfide and ammonia down to lower concentrations. The guard bed 150 maycontain an adsorbent to adsorb impurities such as hydrogen sulfide thatmay poison an oligomerization catalyst. The guard bed 150 may containmultiple adsorbents for adsorbing more than one type of impurity. Atypical adsorbent for adsorbing hydrogen sulfide is ADS-12, foradsorbing CO is ADS-106 and for adsorbing ammonia is UOP MOLSIV 3A allavailable from UOP, LLC. The adsorbents may be mixed in a single bed orcan be arranged in successive beds.

A dilute ethylene stream in line 151 perhaps amine treated, perhapswater washed and perhaps adsorption treated to remove more hydrogensulfide, ammonia and carbon monoxide will typically have at least one ofthe following impurity concentrations: about 0.1 wt-% and up to about5.0 wt-% of carbon monoxide and/or about 0.1 wt-% and up to about 5.0wt-% of carbon dioxide, and/or at least about 1 wppm and up to about 500wppm hydrogen sulfide and/or at least about 1 and up to about 500 wppmammonia, and/or at least about 5 and up to about 20 wt-% hydrogen. Thetype of impurities present and their concentrations will vary dependingon the processing and origin of the dilute ethylene stream.

Line 151 carries the dilute ethylene stream to a compressor 152 to bepressured up to reactor pressure. The compressor 152 is in downstreamcommunication with the main column 92, the product recovery section 90and the product outlet 31. The compressed dilute ethylene stream can becompressed to at least about 3,550 kPa (500 psia) and perhaps no morethan about 10,445 kPa (1500 psia) and suitably between about 4,930 kPa(700 psia) and about 7,687 kPa (1100 psia). It is preferred that thedilute ethylene stream be pressured up to above the critical pressure ofethylene which is about 4,992 kPa (724 psia) for pure ethylene to avoidrapid catalyst deactivation. The compressor 152 may comprise one or morestages with interstage cooling. A heater may be required to bring thecompressed stream up to reaction temperature. The compressed diluteethylene is carried in line 154 to oligomerization reactor 156.

The oligomerization reactor 156 is in downstream communication with thecompressor 152 and the primary and secondary absorbers 114 and 118,respectively. The oligomerization reactor preferably contains a fixedcatalyst bed 158. The dilute ethylene feed stream contacts the catalystpreferably in a down flow operation. However, upflow operation may besuitable. The catalyst is preferably an amorphous silica-alumina basewith a metal from either Group VIII and/or Group VIB in the periodictable using Chemical Abstracts Service notations. In an aspect, thecatalyst has a Group VIII metal promoted with a Group VIB metal. In anaspect, the catalyst has a silica-to-alumina ratio of no more than 30and preferably no more than 20. Typically, the silica and alumina willonly be in the base, so the silica-to-alumina ratio will be the same forthe catalyst as for the base. The metals can either be impregnated ontoor ion exchanged with the silica-alumina base. Co-mulling is alsocontemplated. Catalysts for the present invention may have a LowTemperature Acidity Ratio of at least about 0.15, suitably of about 0.2,and preferably greater than about 0.25, as determined by AmmoniaTemperature Programmed Desorption (Ammonia TPD) as describedhereinafter. Additionally, a suitable catalyst will have a surface areaof between about 50 and about 400 m²/g as determined by nitrogen BET.

A preferred oligomerization catalyst of the present invention isdescribed as follows. The preferred oligomerization catalyst comprisesan amorphous silica-alumina support. One of the components of thecatalyst support utilized in the present invention is alumina. Thealumina may be any of the various hydrous aluminum oxides or aluminagels such as alpha-alumina monohydrate of the boehmite orpseudo-boehmite structure, alpha-alumina trihydrate of the gibbsitestructure, beta-alumina trihydrate of the bayerite structure, and thelike. A particularly preferred alumina is available from Sasol NorthAmerica Alumina Product Group under the trademark Catapal. This materialis an extremely high purity alpha-alumina monohydrate (pseudo-boehmite)which after calcination at a high temperature has been shown to yield ahigh purity gamma-alumina. Another component of the catalyst support isan amorphous silica-alumina. A suitable silica-alumina with asilica-to-alumina ratio of 2.6 is available from CCIC, a subsidiary ofJGC, Japan.

Another component utilized in the preparation of the catalyst utilizedin the present invention is a surfactant. The surfactant is preferablyadmixed with the hereinabove described alumina and the silica-aluminapowders. The resulting admixture of surfactant, alumina andsilica-alumina is then formed, dried and calcined as hereinafterdescribed. The calcination effectively removes by combustion the organiccomponents of the surfactant but only after the surfactant has dutifullyperformed its function in accordance with the present invention. Anysuitable surfactant may be utilized in accordance with the presentinvention. A preferred surfactant is a surfactant selected from a seriesof commercial surfactants sold under the trademark “Antarox” by SolvayS.A. The “Antarox” surfactants are generally characterized as modifiedlinear aliphatic polyethers and are low-foaming biodegradable detergentsand wetting agents.

A suitable silica-alumina mixture is prepared by mixing proportionatevolumes silica-alumina and alumina to achieve the desiredsilica-to-alumina ratio. In an embodiment, 85 wt-% amorphoussilica-alumina with a silica-to-alumina ratio of 2.6 and 15 wt-% aluminapowder will provide a suitable support. In an embodiment, ratios otherthan 85-to-15 of amorphous silica-alumina to alumina may be suitable, solong as the final silica-to-alumina ratio of the support is suitably nomore than 30 and preferably no more than 20.

Any convenient method may be used to incorporate a surfactant with thesilica-alumina and alumina mixture. The surfactant is preferably admixedduring the admixture and formation of the alumina and silica-alumina. Apreferred method is to admix an aqueous solution of the surfactant withthe blend of alumina and silica-alumina before the final formation ofthe support. It is preferred that the surfactant be present in the pasteor dough in an amount from about 0.01 to about 10 wt-% based on theweight of the alumina and silica-alumina.

Monoprotic acid such as nitric acid or formic acid may be added to themixture in aqueous solution to peptize the alumina in the binder.Additional water may be added to the mixture to provide sufficientwetness to constitute a dough with sufficient consistency to be extrudedor spray dried.

The paste or dough may be prepared in the form of shaped particulates,with the preferred method being to extrude the dough mixture of alumina,silica-alumina, surfactant and water through a die having openingstherein of desired size and shape, after which the extruded matter isbroken into extrudates of desired length and dried. A further step ofcalcination may be employed to give added strength to the extrudate.Generally, calcination is conducted in a stream of dry air at atemperature from about 260° C. (500° F.) to about 815° C. (1500° F.).

The extruded particles may have any suitable cross-sectional shape,i.e., symmetrical or asymmetrical, but most often have a symmetricalcross-sectional shape, preferably a spherical, cylindrical or polylobalshape. The cross-sectional diameter of the particles may be as small as40 μm; however, it is usually about 0.635 mm (0.25 inch) to about 12.7mm (0.5 inch), preferably about 0.79 mm ( 1/32 inch) to about 6.35 mm(0.25 inch), and most preferably about 0.06 mm ( 1/24 inch) to about4.23 mm (⅙ inch). Among the preferred catalyst configurations arecross-sectional shapes resembling that of a three-leaf clover, as shown,for example, in FIGS. 8 and 8A of U.S. Pat. No. 4,028,227. Preferredclover-shaped particulates are such that each “leaf” of thecross-section is defined by about a 270° arc of a circle having adiameter between about 0.51 mm (0.02 inch) and 1.27 mm (0.05 inch).Other preferred particulates are those having quadralobalcross-sectional shapes, including asymmetrical shapes, and symmetricalshapes such as in FIG. 10 of U.S. Pat. No. 4,028,227.

Typical characteristics of the amorphous silica-alumina supportsutilized herein are a total pore volume, average pore diameter andsurface area large enough to provide substantial space and area todeposit the active metal components. The total pore volume of thesupport, as measured by conventional mercury porosimeter methods, isusually about 0.2 to about 2.0 cc/gram, preferably about 0.25 to about1.0 cc/gram and most preferably about 0.3 to about 0.9 cc/gram.Ordinarily, the amount of pore volume of the support in pores ofdiameter greater than 100 angstroms is less than about 0.1 cc/gram,preferably less than 0.08 cc/gram, and most preferably less than about0.05 cc/gram. Surface area, as measured by the B.E.T. method, istypically above 50 m²/gram, e.g., above about 200 m²/gram, preferably atleast 250 m²/gram., and most preferably about 300 m²/gram to about 400m²/gram.

To prepare the catalyst, the support material is compounded, as by asingle impregnation or multiple impregnations of a calcined amorphousrefractory oxide support particles, with one or more precursors of atleast one metal component from Group VIII or VIB of the periodic table.The Group VIII metal, preferably nickel, should be present in aconcentration of about 0.5 to about 15 wt-% and the Group VIB metal,preferably tungsten, should be present in a concentration of about 0 toabout 12 wt-%. The impregnation may be accomplished by any method knownin the art, as for example, by spray impregnation wherein a solutioncontaining the metal precursors in dissolved form is sprayed onto thesupport particles. Another method is the multi-dip procedure wherein thesupport material is repeatedly contacted with the impregnating solutionwith or without intermittent drying. Yet other methods involve soakingthe support in a large volume of the impregnation solution orcirculating the support therein, and yet one more method is the porevolume or pore saturation technique wherein support particles areintroduced into an impregnation solution of volume just sufficient tofill the pores of the support. On occasion, the pore saturationtechnique may be modified so as to utilize an impregnation solutionhaving a volume between 10 percent less and 10 percent more than thatwhich will just fill the pores.

If the active metal precursors are incorporated by impregnation, asubsequent or second calcination at elevated temperatures, as forexample, between 399° and 760° C. (750° and 1400° F.), converts themetals to their respective oxide forms. In some cases, calcinations mayfollow each impregnation of individual active metals. A subsequentcalcination yields a catalyst containing the active metals in theirrespective oxide forms.

A preferred oligomerization catalyst of the present invention has anamorphous silica-alumina base impregnated with 0.5-15 wt-% nickel in theform of 3.175 mm (0.125 inch) extrudates and a density of about 0.45 toabout 0.65 g/ml. It is also contemplated that metals can be incorporatedonto the support by other methods such as ion-exchange and co-mulling.

An alternative catalyst suitable for the present invention utilizes aco-gelled silica-alumina support made by the well-known oil-drop methodwhich permits the utilization of the support in the form ofmacrospheres. For example, an alumina sol, utilized as an aluminasource, is commingled with an acidified water glass solution as a silicasource, and the mixture is further commingled with a suitable gellingagent, for example, urea, hexamethylenetetramine, or mixtures thereof.The mixture is discharged while still below gellation temperature, andby means of a nozzle or rotating disk, into a hot oil bath maintained atgellation temperature. The mixture is dispersed into the oil bath asdroplets which form into spheroidal gel particles during passagetherethrough. The alumina sol is preferably prepared by a method whereinaluminum pellets are commingled with a quantity of treated or deionizedwater, with hydrochloric acid being added thereto in a sufficient amountto digest a portion of the aluminum metal and form the desired sol. Asuitable reaction rate is effected at about reflux temperature of themixture.

The spheroidal gel particles prepared by the oil-drop method are aged,usually in the oil bath, for a period of at least 10 to 16 hours, andthen in a suitable alkaline or basic medium for at least 3 to about 10hours, and finally water-washed. Proper gellation of the mixture in theoil bath, as well as subsequent aging of the gel spheres, is not readilyaccomplished below about 48.9° C. (120° F.), and at about 98.9° C. (210°F.), the rapid evolution of the gases tend to rupture and otherwiseweaken the spheres. By maintaining sufficient superatmospheric pressureduring the forming and aging steps in order to maintain water in theliquid phase, a higher temperature can be employed, frequently withimproved results. If the gel particles are aged at superatmosphericpressure, no alkaline aging step is required.

The spheres are water-washed, preferably with water containing a smallamount of ammonium hydroxide and/or ammonium nitrate. After washing, thespheres are dried, at a temperature of from about 93.3° C. (200° F.) toabout 315° C. (600° F.) for a period of from about 6 to about 24 hoursor more, and then calcined at a temperature of from about 426.67° C.(800° F.) to about 760° C. (1400° F.) for a period of from 2 to about 12hours or more.

The Group VIII component and the Group VIB component are composed withthe co-gelled silica-alumina carrier material by any suitableco-impregnation technique. Thus, the carrier material can be soaked,dipped, suspended, or otherwise immersed in an aqueous impregnatingsolution containing a soluble Group VIII salt and a soluble Group VIBsalt. One suitable method comprises immersing the carrier material inthe impregnating solution and evaporating the same to dryness in arotary steam dryer, the concentration of the impregnating solution beingsuch as to ensure a final catalyst composite comprising an atomic ratioof nickel to nickel plus tungsten of about 0.1 to about 0.3. Anothersuitable method comprises dipping the carrier material into the aqueousimpregnating solution at room temperature until complete penetration ofcarrier by the solution is achieved. After absorption of theimpregnating solution, the carrier is drained of free surface liquid anddried in a moving belt calciner.

The catalyst composite is usually dried at a temperature of from about93.3° C. (200° F.) to about 260° C. (500° F.) for a period of from about1 to about 10 hours prior to calcination. In accordance with the presentinvention, calcination is effected in an oxidizing atmosphere at atemperature of from about 371° C. (700° F.) to about 650° C. (1200° F.).The oxidizing atmosphere is suitably air, although other gasescomprising molecular oxygen may be employed.

A suitable alternative catalyst is an oil dropped silica-aluminaspherical support with a diameter of 3.175 mm (0.125 inch) impregnatedwith about 0.5 to about 15 wt-% nickel and with 0 to about 12 wt-%tungsten. Other metals incorporation methods may be suitable and arecontemplated. A suitable density range for the alternative catalystwould be between about 0.60 and about 0.70 g/mL.

The dilute ethylene feed may be contacted with the oligomerizationcatalyst at a temperature between about 200° and about 400° C. Thereaction takes place predominantly in the gas phase at a GHSV 50 to 1000hr⁻¹ on ethylene basis. We have found, surprisingly, that despite thepresence of impurities in the feed that poison the catalyst and dilutionof the ethylene in the feed, that at least about 40 wt-% and as much asabout 75 wt-% of the ethylene in the feed stream convert to heavierhydrocarbons. The ethylene will first oligomerize over the catalyst toheavier olefins. Some of the heavier olefins may cyclize over thecatalyst, and the presence of hydrogen could facilitate conversion ofthe olefins to paraffins which are all heavier hydrocarbons thanethylene.

The catalyst remains stable despite the impure feed, but it can beregenerated upon deactivation. Suitable regeneration conditions includesubjecting the catalyst, for example, in situ, to hot air at 500° C. for3 hours. Activity and selectivity of the regenerated catalyst iscomparable to fresh catalyst.

The oligomerization product stream from the oligomerization reactor inline 160 can be transported to an oligomerization separator 162 whichmay be a simple flash drum to separate a gaseous stream from a liquidstream. The oligomerization separator 162 is in downstream communicationwith the oligomerization reactor 156. The gaseous product stream inoverhead line 164 comprising light gases such as hydrogen, methane,ethane, unreacted olefins and light impurities may be transported to acombustion unit 166 to generate steam in line 167. Alternatively, thegaseous product in overhead line 164 may be combusted to fire a heater(not shown) and/or to provide a source of flue gas to turn a gas turbine(not shown) to generate power. The overhead line 164 is in upstreamcommunication with the combustion unit 166. The liquid bottoms streamcomprising heavier hydrocarbons in line 168 from the oligomerizationseparator 162 can be let down over a valve and recycled back to theproduct separation section 90. The recycle line 168 is in downstreamcommunication with a bottoms 169 of the oligomerization separator 162.Consequently, the main column 92 is in downstream and upstreamcommunication with the oligomerization reactor 156. The bottoms streamis preferably recycled via recycle line 168 to the main column 92 at alocation between the heavy naphtha outlet 96 a and the light cycle oiloutlet 95 a. Alternatively, the recycle line 168 feeds the light cycleoil line 95 or the heavy naphtha line 96. The recycle line is indownstream communication with the oligomerization reactor 156 and inupstream communication with the main column 92. Alternatively, theoligomerization product in lines 160 or 168 may be saturated or not andtransported to a fuel tank without recycling to the product separationzone 90.

EXAMPLES

The utility of the present invention will be demonstrated by thefollowing examples.

Example 1

A nickel and tungsten on an amorphous silica-alumina oil-droppedspherical base was synthesized via the procedures given hereinabove forthe alternative catalyst of the present invention. The metals comprised1.5 wt-% nickel and 11 wt-% tungsten of the catalyst. The sphericalbases had diameters of 3.175 mm. The catalyst had a silica-to-aluminaratio of about 3, a density of 0.641 g/mL and a surface area of 371m²/g.

Example 2

An extruded amorphous silica-alumina was synthesized by combining anamorphous silica-alumina having a silica-to-alumina ratio of about 2.6provided by CCIC, and pseudo-boehmite provided under the Catapaltrademark in a weight ratio of 85-to-15. The pseudo-boehmite waspeptized with nitric acid before mixture with the amorphoussilica-alumina. A surfactant provided under the Antarox trademark andwater in sufficient quantity to wet the dough were added to the mixture.The catalyst dough was extruded through 1.59 mm openings in acylindrical die plate and broken into pieces prior to calcination at550° C. The finished catalyst consisted of 85 wt-% silica-alumina and 15wt-% alumina, had a silica-to-alumina ratio of 1.92 and had a surfacearea of 368 m²/g.

Example 3

Of Ni(NO₃)₂.6H₂O, 3.37 grams was dissolved in 32.08 grams of deionizedwater. The nickel solution was contacted with the extruded amorphoussilica-alumina of Example 2 by adding the nickel solution in fourths andshaking vigorously between additions. A light green extrudate resulted.The nickel metal was then converted to the oxide form by drying theextrudates at 110° C. for 3 hours, then calcining by ramping to 500° C.at 2° C./min and holding at 500° C. for 3 hours before cooling to roomtemperature. The light gray extrudates were found to contain 1.5 wt-%nickel.

Example 4

A sample of MTT zeolite with a silica-to-alumina ratio of 40 wasobtained from the Zeolyst Corporation. The MTT zeolite was combined withpseudo-boehmite and extruded through 3.175 mm openings in a cylindricaldie plate before calcining to 550° C. The finished catalyst consisted of80 wt-% MTT zeolite and 20 wt-% alumina.

Example 5

The catalyst of Example 1 was tested for olefin oligomerization at 280°C., 6,895 kPa (1000 psig), 58 OGHSV (olefin gas hourly space velocity)in a fixed bed operation over 10 mL of catalyst. The feed consisted of30 wt-% C₂H₄ and 70 wt-% CH₄. Results are shown in Table I.

Example 6

The catalyst of Example 2 was tested for olefin oligomerization at 280°C., 6,895 kPa (1000 psig), 586 OGHSV in a fixed bed operation over 10 mLof catalyst. The feed consisted of 23 wt-% C₂H₄, 14 wt-% C₂H₆, 35 wt-%CH₄, 13 wt-% H₂, 13 wt-% N₂, 1 wt-% CO, 1.5 wt-% CO₂, 10 wppm H₂S andwas saturated with water vapor at 25° C. and 3,447 kPa (500 psig) priorto feeding the oligomerization reaction. Results are shown in the TableI and in FIG. 2.

Example 7

The catalyst of Example 3 was tested for olefin oligomerization at 280°C., 6,895 kPa (1000 psig), 586 OGHSV in a fixed bed operation over 10 mLof catalyst. The feed consisted of 23 wt-% C₂H₄, 14 wt-% C₂H₆, 35 wt-%CH₄, 13 wt-% H₂, 13 wt-% N₂, 1 wt-% CO, 1.5 wt-% CO₂, 10 wppm H₂S andwas saturated with water vapor at 25° C. and 3,447 kPa (500 psig) priorto feeding the oligomerization reaction. Results are shown in Table Iand FIG. 2. During 27-44 hours on stream, 1 ppm NH₃ was also added tothe feed. No changes in conversion or selectivity were noted.

Example 8

The experiment of Example 7 was repeated except that the concentrationof H₂S in the feed was 50 wppm rather than 10 wppm. Results are shown inTable I and in FIG. 2.

FIG. 2 is a plot of C₂H₄ conversion versus time on stream for Examples6-8. The nickel on amorphous silica-alumina catalyst of Example 3performed better than just the silica-alumina base of Example 2 in termsof ethylene conversion. The catalysts of Examples 2 and 3 were alsolittle affected by feed impurities.

Example 9

The catalyst of Example 4 was tested for olefin oligomerization at 280°C., 6,895 kPa (1000 psig), 586 OGHSV in a fixed bed operation over 10 mLof catalyst. The feed consisted of 23 wt-% C₂H₄, 14 wt-% C₂H₆, 35 wt-%CH₄, 13 wt-% H₂, 13 wt-% N₂, 1 wt-% CO, 1.5 wt-% CO₂, 10 wppm H₂S andwas saturated with water vapor at 25° C. and 3,447 kPa (500 psig) priorto feeding the oligomerization reaction. Results are shown in Table Iand in FIG. 3.

FIG. 3 is a plot of C₂H₄ conversion versus time on stream for Example 9showing the effect of impurities that can poison the MTT zeolitecatalyst of Example 4. After 20 hours of reaction, conversion haddropped to below 10 wt-%, while conversion by the catalyst from Example3 in Examples 7 and 8 were maintained around 60 wt-%.

Example 10

The catalyst of example 4 was tested for olefin oligomerization at 280°C., 6,895 kPa (1000 psig)kPa, 613 OGHSV in a fixed bed operation over 10mL catalyst. The feed consisted of 30 wt-% C₂H₄ and 70 wt-% CH₄. At 21hours on stream, hydrogen was added to achieve a feed consisting of 27wt-% C₂H₄, 63 wt-% CH₄ and 10 wt-% H₂. At 45 hours on stream, 500 wppmNH₃ in H₂ was added to achieve a feed of 27 wt-% C₂H₄, 63 wt-% CH₄, 10wt-% H₂ and 50 wppm NH₃. Results are shown in Table I and in FIG. 4.

FIG. 4 is a plot of C₂H₄ conversion versus time on stream from Example10 showing the effect of impurities H₂ and NH₃ on the MTT zeolitecatalyst of Example 4. As can be seen in FIG. 3, ethylene conversiondropped upon introduction of hydrogen to the feed at 20 hours on stream.Additionally, upon introduction of the ammonia at 45 hours on stream,ethylene conversion quickly decreased significantly.

TABLE I C₂-C₄ C₅-C₁₀ C₁₀ ⁺ Example Conversion Selectivity SelectivitySelectivity Example 5 92 30 53 17 Example 6 38 8 34 58 Example 7 75 9 1873 Example 8 74 9 17 74 Example 9  5 hours on stream 77 1 31 68 20 hourson stream 7 10 35 55 Example 10 10 hours on stream 90 1 31 68 35 hourson stream 65 4 31 65 45 hours on stream 63 4 31 65 70 hours on stream 306 24 70

Example 11

The Ammonia Temperature Programmed Desorption (Ammonia TPD) testinvolves first heating about a 250 milligram sample of catalyst at arate of about 5° C. per minute to a temperature of about 550° C. in thepresence of 20 volume percent oxygen in helium atmosphere at a flow rateof about 100 milliliters per minute. After a hold of about one hour,helium is used to flush the system for about 15 minutes and the sampleis cooled to about 150° C. The sample is then saturated with pulses ofammonia in helium at about 40 milliliters per minute. The total amountof ammonia used is greatly in excess of the amount required to saturateall the acid sites on the sample. The sample is purged with helium atabout 40 milliliters per minute for about 8 hours to remove physisorbedammonia. With the helium purge continuing, the temperature is increasedat a rate of about 10° C. per minute to a final temperature of 600° C.The amount of ammonia desorbed is monitored using a calibrated thermalconductivity detector. The total amount of ammonia is determined byintegration.

The ratio of the total amount of ammonia desorbed to the dry weight ofthe sample yields the Total Acidity. As used herein, values of TotalAcidity are given in units of millimoles of ammonia per gram of drysample. Catalysts active for the oligomerization of dilute ethylenestreams are acidic, that is, having a Total Acidity of at least about0.15, and preferably at least about 0.25, as determined by Ammonia TPD.

The ratio of the total amount of ammonia desorbed from the sample priorto reaching a temperature of 300° C. to the dry weight of the sampleyields the Low Temperature Peak. As used herein, values of the LowTemperature Peak are given in units of millimoles of ammonia per gram ofdry sample. Catalysts active for the oligomerization of dilute ethylenestreams have a Low Temperature Peak, that is having a Low TemperaturePeak of at least 0.05, and preferably at least 0.06, as determined byAmmonia TPD.

The ratio of the Low Temperature Peak to the Total Acidity gives aunit-less ratio known as Low Temperature Acidity Ratio. Catalystsresistant to feed impurities in dilute ethylene streams active for theoligomerization of dilute ethylene streams have a Low TemperatureAcidity Ratio of at least 0.15, suitably at least 0.2, and preferablygreater than 0.25, as determined by Ammonia TPD.

TABLE II Total Acidity, Low Temperature Peak, Low Temperature Catalystmillimoles/g millimoles/g Acidity Ratio Example 1 0.323 0.116 0.36Example 2 0.285 0.092 0.32 Example 3 0.264 0.084 0.32 Example 4 0.3110.037 0.12

It can be seen from the Examples that the zeolitic catalyst is renderedmuch less effective in terms of ethylene conversion by feed impuritieswhile the catalysts of the present invention remain an effectiveethylene oligomerization catalyst despite the presence of impurities inthe feed which are typical catalyst poisons. The catalysts of thepresent invention maintain ethylene conversions of at least 40 wt-%,typically 60 wt-% and preferably above 70 wt-%.

Without further elaboration, it is believed that one skilled in the artcan, using the preceding description, utilize the present invention toits fullest extent. The preceding preferred specific embodiments are,therefore, to be construed as merely illustrative, and not limitative ofthe remainder of the disclosure in any way whatsoever.

In the foregoing, all temperatures are set forth in degrees Celsius and,all parts and percentages are by weight, unless otherwise indicated.

From the foregoing description, one skilled in the art can easilyascertain the essential characteristics of this invention and, withoutdeparting from the spirit and scope thereof, can make various changesand modifications of the invention to adapt it to various usages andconditions.

1. A process comprising: contacting cracking catalyst with a hydrocarbon feed stream to crack the hydrocarbons to cracked product hydrocarbons having lower molecular weight and deposit coke on the cracking catalyst to provide coked cracking catalyst; separating said coked cracking catalyst from said cracked products; adding oxygen to said coked cracking catalyst; combusting coke on said coked cracking catalyst with oxygen to regenerate said cracking catalyst; separating said cracked products to obtain a dilute ethylene stream comprising between about 5 and about 50 wt-% ethylene and between about 25 and about 55 wt-% methane; compressing the dilute ethylene stream to a pressure of between about 4,826 and about 7,584 kPa; and contacting the dilute ethylene stream with a oligomerization catalyst.
 2. The process of claim 1 further comprising converting at least 40 wt-% of the ethylene in the feed stream to heavier hydrocarbons.
 3. The process of claim 1 wherein said oligomerization catalyst is an amorphous silica-alumina base and a metal selected from the group consisting of Group VIII and Group VIB in the periodic table.
 4. The process of claim 1 wherein said oligomerization catalyst is an amorphous silica-alumina base impregnated with 0.5 to 15 wt-% nickel.
 5. The process of claim 1 wherein said dilute ethylene stream is compressed to a pressure above the critical pressure for ethylene.
 6. The process of claim 1 wherein said catalyst has a Low Temperature Acidity Ratio of at least about 0.15 as determined by an Ammonia Temperature Programmed Desorption test.
 7. The process of claim 1 wherein said dilute ethylene stream comprises no more than about 0.5 wt-% propylene.
 8. The process of claim 1 wherein said contacting step is performed in a fixed bed of said oligomerization catalyst.
 9. The process of claim 1 further comprising treating said dilute ethylene stream in a guard bed to remove one or more of the impurities prior to said second contacting step.
 10. The process of claim 1 further comprising separating an oligomerization product stream from said contacting step into a gaseous stream and a liquid oligomerization product stream.
 11. The process of claim 10 further comprising recycling said liquid oligomerization product stream to said second separation step.
 12. The process of claim 1 wherein said contacting step is performed at a temperature between about 200° C. and about 400° C.
 13. The process of claim 1 wherein said feed is predominantly in the gas phase.
 14. The process of claim 13 wherein said second contacting step is performed at GHSV 50 to 1000 hr⁻¹ on an ethylene basis.
 15. The process of claim 1 wherein said dilute ethylene stream comprises between about 10 and about 30 wt-% ethylene.
 16. The process of claim 1 wherein said dilute ethylene stream further comprises between about 5 and about 45 wt % ethane.
 17. A process comprising: contacting cracking catalyst with a hydrocarbon feed stream to crack the hydrocarbons to cracked product hydrocarbons having lower molecular weight and deposit coke on the cracking catalyst to provide coked cracking catalyst; separating said coked cracking catalyst from said cracked products; adding oxygen to said coked cracking catalyst; combusting coke on said coked cracking catalyst with oxygen to regenerate said cracking catalyst; separating said cracked products to obtain a dilute ethylene stream comprising between about 25 and about 55 wt-% methane, between about 5 and about 50 wt-% ethylene and no more than about 0.5 wt-% propylene; contacting the dilute ethylene stream with a oligomerization catalyst comprising an amorphous silica-alumina base and a metal selected from the group consisting of Group VIII and Group VIB in the periodic table; and converting at least 40 wt-% of the ethylene in the dilute ethylene stream to heavier hydrocarbons.
 18. The process of claim 17 further comprising separating an oligomerization product stream from said contacting step into a gaseous stream and a liquid oligomerization product stream and recycling said liquid oligomerization product stream to said second separation step.
 19. The process of claim 17 further compressing the dilute ethylene stream to a pressure of at least 3,447 kPa.
 20. The process of claim 17 wherein said contacting step is performed in a fixed bed of said oligomerization catalyst. 